Catalytic conversion of hydrocarbon mixtures containing asphaltenes



United States Patent Olfice 3,530,062 Patented Sept. 22, 1970 3,530,062 CATALYTIC CONVERSION OF HYDROCARBON MIXTURES CONTAINING ASPHALTENES John G. Gatsis, Des Plaines, Ill, assignor to Universal Oil Products Company, Des Plaines, 111., a corporation of Delaware No Drawing. Filed May 19, 1967, Ser. No. 639,713 Int. Cl. (110g 23/00 US. Cl. 208-89 3 Claims ABSTRACT OF THE DISCLOSURE Conversion of asphaltene-containing hydrocarbonaceous mixtures commonly referred to as black oils. Applicable principally where the charge stock is additionally contaminated by sulfurous and/or nitrogeneous compounds as well as organo-metallic complexes. The charge stock is initially processed at a catalyst temperature below 430 C. (806 F.) and subsequently at a maximum catalyst temperature above 430 C.

APPLICABILITY OF INVENTION The invention herein described involves a multiplestage process for the hydrorefining and conversion of heavy hydrocarbon fractions and/or distillates, and is especially advantageous in treating petroleum crude oils and topped, or reduced crude oils for the removal of organo-metallic contaminants, nitrogenous and sulfurous compounds, and the conversion of heptane-insoluble asphaltenic material.

Petroleum crude oils, and topped or reduced crude oils, as well as other heavy hydrocarbon fractions and/or distillates, including black oils, heavy cycle stocks, visbreaker liquid efiluent, crude tower bottoms product, tar sand oils, etc., are contaminated by the inclusion of excessive quantities of various non-metallic and metallic impurities. Among the non-metallic impurities are nitrogen, sulfur and oxygen which exist as heteroatomic compounds. Both nitrogenous and sulfurous compounds are objectionable since the combustion of fuels containing these impurities results in the release of nitrogen oxides and sulfur oxides, presenting a serious problem with respect to atmospheric pollution.

In addition to the foregoing described contaminating influences, petroleum crude oils and other heavy hydrocarbonaceous material contain high molecular weight asphaltenic compounds. These are non-distillable, oil-insoluble coke precursors which may be complexed with sulfur, nitrogen, oxygen and various metals. Although the metallic contaminants may exist within the hydrocarbonaceous material in a variety of forms, they are generally present as organo-metallic compounds of relatively high molecular weight, such as metallic porphyrins. A considerable quantity of the organo-metallic complexes are linked with asphaltenic material and become concentrated in the residual fraction; other organo-metallic complexes are volatile, oil-soluble and are, therefore, present in the lighter distillate fraction-Le. boiling below about 1,050 F. (621 C.). A reduction in the concentration of the organometallic complexes is not easily achieved, and to the extent that the crude oil, reduced crude oil, or other heavy hydrocarbon charge stock derived therefrom becomes suitable for further processing. Notwithstanding that the concentration of these organo-metallic complexes may be relatively small in distillate oils, for example, often less than about ppm, calculated as if the complex existed as the elemental metal, subsequent processing techniques are adversely affected thereby. With respect to a process for hydrorefining or treating of hydrocarbon fractions and/ or distillates, the presence of large quantities of asphaltenic material and organo-metallic compounds interferes considerably with the activity of the catalyst with respect to the destructive removal of the nitrogenous, sulfurous and oxygenated compounds, which function is normally the easiest for the catalytic composite to perform to an acceptable degree.

A wide variety of heavy hydrocarbon fractions and/or distillates may be converted and treated, or decontaminated effectively through the utilization of the method of the present invention. Such heavy hydrocarbon fractions include full boiling range crude oils, topped or reduced crude oils, atmospheric distillates, visbreaker bottoms product, heavy cycle stock from thermally or catalytically-cracked charge stocks, heavy vacuum gas oils, shale oil, tar sand oils, etc. A Wyoming sour crude oil, having a gravity of 232 API at 60 F., is contaminated by the presence of 2.8% by weight of sulfur, 2,700 ppm, of total nitrogen, approximately ppm. of metallic complexes, computed as elemental metals, and contains a high boiling, insoluble asphaltenic fraction in an amount of about 8.5% by weight. A more difficult charge stock to convert into useful liquid hydrocarbons, is a crude tower bottoms product, having a gravity, degrees API at 60 F., of 14.3, and contaminated by the presence of 3.0% by weight of sulfur, 3,830 ppm. of total nitrogen, ppm. of total metals and about 10.93% by weight of asphaltenic product, having a gravity of 70 API at 60 F., is contaminated by 6,060 ppm. of total metals, and contains an asphaltenic fraction in an amount of 24.0% by weight.

PRIOR ART It must be acknowledged that published literature recognizes various types of processes designed to effect the hydrorefining and conversion of black oils. Thus, many literature references and/or publications might be found which disclose propane deasphalting followed by thermal cracking or coking of the resulting normally liquid product, desalting followed by halogen hydride treatment to coagulate the metallic-containing asphaltenes, etc. It is noteworthy that the latter processing schemes are unconcerned with catalytic processing of black oils.

Furthermore, with respect to catalytic processing, two principal approaches have been advanced: liquid-phase hydrogenation and vapor-phase hydrocracking. In the former, liquid-phase oil is passed (generally upwardly), in admixture with hydrogen, into a fixed-fluidized bed of catalyst particles. Although perhaps effective in converting at least a portion of the oil-soluble metallic complexes, this type of process is relatively ineffective with respect to asphaltness dispersed in the charge, since the probability of effecting simultaneous contact between the catalyst particle and the asphaltenic molecule is remote. Furthermore, since the reaction zone is generally maintained at an elevated temperature of at least about 500 C. (932 F.), the retention of unconverted asphaltenes suspended in a free liquid phase oil for an extended period of time, results in further agglomeration, making conversion thereof substantially more difficult. Some processes have been described which rely extensively upon thermal cracking in the presence of hydrogen; such processes are unable to effect the conversion of asphaltenic material. In the preferred catalytic process, in which the asphaltenic material is maintained in a dispersed state in a liquid phase rich in hydrogen, the material comes into intimate contact with a fixed-bed of catalyst capable of effecting reaction between the hydogen and the asphaltenes; the liquid phase is itself dispersed in a hydrogen-rich gas phase so the dissolved hydrogen is continuously being replenished. The two-fold dispersion and rapid, intimate contact with the catalytic surface overcomes the difficulties encountered in previous processes whereby excessive residence times and depletion of localized hydrogen supply permit agglomeration of asphaltenes and other high molecular weight species. Such agglomerates are even less available to hydrogen and are not, therefore susceptible to catalytic reaction. They eventually form coke which becomes deposited on the catalyst, thereby further reducing catalytic activity within the system.

Thus, while it is seen that the preferred method of the prior art involves the use of a fixed-bed catalytic technique, there are certain difficulties which continue to plague an efficient, economical operation. While these difficulties are partially solved by a moving-bed or slurry operation, such a process tends to result in a high degree of erosion, thereby causing plant maintenance and replacement of process equipment to be diflicult and expensive. Furthermore, a slurry operation has the disadvantage of having a relatively small amount of catalyst be ing admixed with relatively large quantities of asphaltenic unaterial. In other words, too few catalytically active sites are made available for immediate reaction, with the result that the asphaltenic material has the tendency to undergo thermal cracking, resulting in large quantities of light gases and coke. These difliculties are in turn at least partially avoided through the utilization of a fixed-fluidized process in which the catalytic composite is disposed within a confined reaction zone, being maintained, however, in a fluidized state by exceedingly large quantities of a fast-flowing hydrogen-containing gast stream. Difiiculties attendant the fixed-fluidized bed process reside in a large loss of catalyst, removed from the reaction zone with the hydrocarbon product efiluent, the relatively large quantities of catalyst necessary to effect proper contact between the asphaltenic material and active catalyst sites, etc.

A principal object of my invention is, therefore, to provide a fixed-bed process for the conversion of asphaltene-containing hydrocarbon charge stocks. A corollary objective is to avoid the difiiculties usually attendant a fixed-bed catalytic process while simultaneously taking advantage of the benefits accruing over and above other types of catalytic processing.

Another object involves providing a multiple-stage process which affords a greater degree of asphaltene-conversion to distillable hydrocarbons, thereby increasing the volumetric yield of more valuable hydrocarbons.

These, and other objectives and advantages are achieved through the use of a broad embodiment of the present invention which encompasses a process for the conversion of an asphaltene-containing hydrocarbon charge stock, which process comprises the steps of: (a) admixing said charge stock with hydrogen and catalytically reacting the mixture in a first reaction zone at a maximum catalyst temperature below about 430 C.; (b) further reacting at least a portion of the resulting first reaction zone effluent containing unconverted asphaltenes in a second catalytic reaction zone at a maximum catalyst temperature below about 430 C.; (c) separating the resulting second zone effluent to concentrate unconverted asphaltenes in a residuum fraction; (d) combining the remainder of said second zone efiiuent with at least a second portion of the remainder of said first zone efiluent; and, (e) catalytically reacting the resulting mixture in a third reaction zone at a maximum catalyst temperature above about 430 C.

Other embodiments of my invention reside in the use of particular operating conditions and internal recycle streams. Thus, although the maximum catalyst temperature within the first reaction zone is below 43 C. (806 F.), the lower limit thereof is 320 C. (608 F.). Similarly, with respect to the second reaction zone, the maximum catalyst temperature is within the range of from 320 C. to about 430 C. An elevated maximum catalyst temperature is utilized in the third reaction zone, and is in the range of 430 C. to about 510 C. (950 F.). In a particularly preferred embodiment, the reaction product eflluent from the first reaction zone is separated at a cut-point of about 454 C. (850 F.) to provide an 850 F.-plus heavier fraction containing those asphaltenes unconverted in the first reaction and which are then subjected to conversion in the second reaction zone. At least a portion of this heavier fraction is recycled to combine with the fresh hydrocarbon charge stock to serve as a diluent therefor. The effluent from the second reaction zone is separated at a substantially reduced pressure less than about p.s.i.g. in order to concentrate the remaining unconverted asphaltenes in the residuum fraction.

SUMMARY OF INVENTION The inventive concept encompassed by the foregoing described embodiments stems from recognition of the character of native asphaltenes found in the crude oil and heavier fractions thereof, and the character of residual asphaltenes after the crude oil has been precessed catalytically in the presence of hydrogen. Statistical constitutional analyses, spectroscopic and chemical studies indicate that the native petroleum asphaltene fraction consists of molecules having a similar structure, but having a relatively wide molecular weigh trange. The native asphaltenes, prior to processing, are shown to have a hydrogenzcarbon atomic ratio of about 1.0. When these native asphaltenes are processed at a temperature exceeding 430 C., the hydrogenzcarbon ratio of the residual, unconverted asphaltenes is less than 1.0, causing further conversion to be extremely diflicult if not in fact economically impossible. Analyses have further indicated that the residual asphaltenes, when processing the charge stock at a temperature below about 430 C., are not appreciably different from the native asphaltenes, and have a hydrogenzcarbon atomic ratio of about 1.0. Such residual asphaltenes are not exceptionally refractory with respect to further processing. As hereinafter indicated by a specific example, with a charge stock previously processed at below 430 C., 87.5% of the residual asphaltenes were converted during subsequent procesing, whereas only about 43.1% conversion of residual asphaltenes was effected when the charge stock had been previously processed above 430 C.

In accordance with the present invention, this concept is advantageously employed in the multiple-stage processing of black oils containing asphaltenes in addition to other contaminating influences. Before describing my invention in greater detail, several definitions are believed desirable in order that clear understanding be afforded. In the present specification, and where used in the appended claims, the phrase, temperature substantially the same as, is employed to indicate that the only reduction in temperature stems from normally experienced loss due to the flow of material from one piece of equipment to another, or from conversion of sensible to latent heat by flashing where a pressure drop occurs. The phrase, hydrocarbons boiling within the gasoline boiling range, or gasoline boiling range hydrocarbons, is intended to connote those hydrocarbons boiling at temperatures up to about 400 F. including Q -hydrocarbons, and, as in some localities, "C -hydrocarbons. Although the end boiling point of gasoline is sometimes considered to be as high as 425 F., or even 450 F., the use of the term herein will allude to a hydrocarbon fraction having a nominal end boiling point of about 400 F. Middle-distillate hydrocarbons, whether inclusive of kerosene, will connote a hydrocarbon mixture boiling above 400 F. and having a nominal end boiling point in the range of 650 F. to about 700 F.

Likewise, a black oil is intended to connote a hydrocarbonaceous mixture of which at least about 10.0% boils above a temperature of about 1,050 E, and which has gravity degrees API at 60 F., of 20.0 or less. Distillable hydrocarbons are those normally liquid hydrocarbons, including pentanes, having boiling points below about 1,050 F. Conversion conditions are intended to be those conditions imposed upon the conversion zones in order to convert a substantial portion of the black oil to distillable hydrocarbons. Since the bulk of the reactions being effected are exothermic, the temperature increases through the catalyst bed, and the reaction zone eflluent will be at a temperature higher than at the inlet to the catalyst bed. Hydrogen is mixed with the charge stock, by means of compressive recycle, in an amount of from about 1,000 to about 50,000 s.c.f./bbl., at the selected operating pressure, and preferably in an amount of from about 1,000 to about 10,000 s.c.f./bbl. The operating pressure will be greater than 1,000 p.s.i.g., and generally in the range of about 1,500 p.s.i.g. to about 4,000 p.s.i.g. The black oil passes through the catalyst at a liquid hourly space velocity (defined as volumes of liquid hydrocarbon charge per hour, measured at 60 F. per volume of catalyst disposed in the reaction zone) of from about 0.25 to about 5.0. It is particularly preferred to introduce the mixture of black oil and hydrogen into the reaction vessels in such a manner that the same passes through in downward flow. The internals of the various reaction vessels may be constructed in any suitable manner capable of providing the required intimate contact between the liquid charge stock, the gaseous mixture and the catalyst. In some instances, it may be desirable to provide the reaction zone with a packed bed of inert material such as particles of granite, porcelain, Berl saddles, sand, aluminum or other metal turnings, etc., to facilitate distribution of the charge, or to employ perforated trays or special mechanical devices for this purpose.

The temperature of the charge mixture to the first reaction zone is controlled such that the maximum catalyst temperature within the zone is in the range of from about 320 C. to about 430 C. The reaction product efiluent is then separated at a cut point of about 440 C. to about 460 C. to provide a heavier fraction containing the unconverted asphaltenes. In a preferred operation, a portion of this heavier fraction is recycled to combine with the fresh charge stock and hydrogen in an amount such that the combined liquid feed ratio is in the range of about 1.211 to about 2.5 :1. The remainder of the heavier fraction is subjected to further reaction in a second reaction zone at about the same conditions of temperature and pressure as the first reaction zone with, however, a greater hydrogen concentrationi.e. from about 1,000 to about 3,000 s.c.f./bbl. more than the first zone. The effluent from the second reaction Zone is hot flashed in a separation zone maintained at less than about 100* p.s.i.g., at substantially the same temperature existing at the outlet of the second reaction zone. The hot flash separation may consist of an initial low pressure flash chamber maintained at about 60 p.s.i.g., in combination with a vacuum column maintained at about 5060 mm. of Hg absolute. This hot flash system serves to concentrate the remaining unconverted asphaltenes in a residuum fraction containing a significant amount of the sulfurous compounds not converted in the preceding reaction zones, and to flash the distillable hydrocarbons. This flashed material, having an end boiling point of about 1,050 F is combined with the light fraction separated from the first reaction zone effluent, and, subjected to further conversion dependent upon the ultimate desired product distributione.g. whether to maximize gasoline boiling range hydrocarbons, kerosene or middle-distillate fractions.

The catalytic composite disposed within the conversion zones can be characterized as comprising a metallic component having hydrogenation activity, and composited with a refractory inorganic oxide carrier material of either synthetic or natural origin. The precise composition and method of manufacturing the carrier material is not considered essential to the present process, although a siliceous carrier, such as 88.0% alumina and 12.0% silica, or 63.0% alumina and 37.0% silica, are generally preferred. Suitable metallic components having hydrogenation activity are those selected from the group consisting of the metals of Groups VI-B and VIII of the Periodic Table, as indicated in the Periodic Chart of the Elements, Fisher Scientific Company (1953). Thus, the catalytic composite may comprise one or more metallic components from the group of molybdenum, tungsten, chromium, iron, cobalt, nickel, platinum, palladium, iridium, osmium, rhodium, ruthenium, and mixtures thereof. The concentration of the catalytically active metallic compo nent, or components, is primarily dependent upon the particular metal as well as the characteristics of the charge stock. For example, the metallic components of Group VI-B are preferably present in an amount within the range of about 1.0% to about 20.0% by weight, the iron-group metals in an amount within the range of about 0.2% to about 10.0% by weight, whereas the platinumgroup metals are preferably present in an amount within the range of about 0.1% to about 5.0% by weight, all of which are calculated as if the components existed within the finished catalytic composite as the elemental metal.

The refractory inorganic oxide carrier material may comprise alumina, silica, zirconia, magnesia, titania, boria, strontia, hafnia, and mixtures of two or more including silica-alumina, alumina-silica-boron phosphate, silicazirconia, silica-magnesia, silica-titania, alumna-zirconia, alumina-magnesia, alumina-titania, magnesia-zirconia, titania-zirconia, magnesia-titania, silica-alumina-zirconia, silica-alumina-magnesia, silica-alumina-titania, silica-magnesia-zirconia, silica-alumina-boria, etc. It is preferred to utilize a carrier material containing at least a portion of silica, and preferably a composite of alumina, silica and boron phosphate with alumina being in the greater proportion. The function of these first reaction zones is essentially two-fold; they serve to concentrate a residuum fraction containing sulfur while simultaneously producing distillable hydrocarbons. The quantity of residual sulfur is dependent upon the characteristics of the fresh black oil charge stock. However, most of this remaining sulfur is concentrated in the middle-distillate and gas oil ranges-from about 525 F. to about 1,050" F. along with any residual nitrogenous compounds.

The last in the series of reaction zones is utilized primarily to effect the virtually complete removal of nitrogenous and sulfurous compounds from the distillables, and to produce lower boiling gasoline and middle-distillate hydrocarbons boiling up to about 850 F. The ranges of variable operating conditions in the final reaction zone include temperatures, as measured at the inlet to the catalyst bed, of from about 500 F. to about 850 F., pressures of 1,000 to about 4,000 p.s.i.g. hydrogen concentrations of from about 1,000 to about 50,000 s.c.f./bbl. and liquid hourly space velocities of from about 0.5 to about 5.0. Precise values for any one or more of these variables depends essentially upon two aspects which are necessarily considered in a given situation: (1) the nature of the distillable hydrocarbon mixture resulting from the first and second reaction zones, and (2) the ultimately desired product distribution-Le. the relative proportions of gasoline, kerosene and middle-distillate boiling range hydrocarbons.

The effluent from the hydrocracking zone is introduced into a cold high-pressure separator functioning at substantially the same pressure as the outlet of the conversion zone, and at a substantially lower temperature in the range of 60 F. to about F. The gaseous phase, rich in hydrogen, is recycled at least in part to this hydrocracking zone, and at least in part with the charge to both the first and second reaction zones. The liquid phase from the cold separator constitutes the product of the present invention, and may be subjected to standard fractionation and other separation techniques for the purpose of recovering specific select fractions.

It is preferred that the catalytic composite in the hydrocracking zone comprise at least two refractory inorganic oxides, and preferably alumina and silica. When silica and alumina are employed in combination, the lat ter will be present within an amount of from about 10% to about 90% by weight. Excellent results have been achieved through the utilization of the following silicaalumina composites: 88% by weight of silica and 12% by weight of alumina, 63% by weight of alumina and 37% by weight of silica, 88% by Weight alumina and 12% by weight of silica. Suitable catalytic composites, for utilization in the hydrocracking zone, comprise the following, but not by way of limitation: 6.0% by weight of nickel and 0.2% by weight of molybdenum; 1.8% by weight of nickel and 16.0% by weight of molybdenum; 6.0% by weight of nickel; 0.4% by weight of palladium; 6.0% by weight of nickel and 0.2% by weight of platinum; 6.0% by weight of nickel and 0.2% by weight of iron; 0.4% by weight of platinum; 6.0% by weight of nickel; 12.0% by weight of molybdenum; and 6.0% by weight of nickel and 0.2% by weight of palladium. In many instances, particularly when the catalytically active metallic components comprise metals selected from the platinum-group of Group VIII of the Periodic Table it will be desirable to include a halogen component to impart an additional acid-acting function to the hydrocracking catalyst. It is understood that the broad scope of the present invention is not to be unduly limited to the utilization of a particular catalyst having a particular concentration of components, a particular means for the manufacture of the same, or specific operating conditions other than those previously set forth. The utilization of any of the previously mentioned catalytic composites, at operating conditions varying within the limits hereinbefore set forth does not necessarily yield results equivalent to the utilization of the other catalytic composites employed under other operating conditions.

The gaseous ammonia and hydrogen sulfide, resulting from the destructive removal of nitrogenous and sulfurous compounds, and light paraffinic hydrocarbons, are removed from the total efiluent in any suitable manner. For example, the efiluent may be admixed with water, and thereafter subjected to separation such that the ammonia is absorbed in the water-phase. Hydrogen sulfide and light parafiinic hydrocarbons may be removed by introducing the effluent into a low-temperature flash chamber, the normally liquid hydrocarbons from which are passed into a fractionating column for the purpose of removing those hydrocarbons boiling within the gasoline boiling range. Other conditions and preferred operating techniques will be presented in conjunction with the following examples, one of which refers to the applicability of my invention in a commercially-scaled unit.

EXAMPLES Recognition of the concept encompassed by the present invention was facilitated during the processing of the crude tower bottom obtained from a sour Wyoming crude. The analysis of this charge stock is presented in Table I.

The asphaltenes in this charge stock have an average molecular weight of about 4,580, and contain 84.15% carbon, 7.23% hydrogen, 1.72% nitrogen and 4.54% sulfur. The atomic ratio of hydrogenzcarbon is 1.02. In an operation at a pressure of 3,000 p.s.i.g., a liquid hourly space velocity of 4.0, and a temperature gradient of 370 to 400 C. (entrance and exit of the catalyst bed), the heptane-insoluble portion decreased to 6.2% by weight 8 and 57.5 volume percent was distillable at 1,000" F. Analyses performed on the asphaltenes indicated little change in the character thereof from the native asphaltenes; they possessed an average molecular weight of 4,566 and contained 4.38% sulfur, 1.21% nitrogen, 84.01% carbon and 7.30% hydrogen, the hydrogenzcarbon atomic ratio being 1.03.

With a temperature gradient of 380 to 425 C., all else remaining the same, the total product contained 2.9% heptane-insolubles with 63.0% by volume being distillable at 1,000 F. The residual asphaltenes appear to have the same structural group, but a lower molecular weight of 3,962. Analysis indicates 2.75% sulfur, 1.20% nitrogen, 87.81% carbon and 7.26% hydrogen for a hydrogenzcarbon atomic ratio of 0.98.

The asphaltenes separated tom the liquid product resulting from a temperature gradient of 420 to 470 C., containing 1.8% heptane-insolubles, and having 93.0%- distillable at 1,000 F., had an average molecular weight of 1,208. They consist of 1.05% sulfur, 1.30% nitrogen, 91.16% carbon and 5.93% hydrogen for a hydrogen:- carbon atomic ratio of 0.77. The lower atomic ratio of hydrogenzcarbon indicates a dissimilar structural group from the native asphaltenes and one which is characterized by a more highly condensed ring system. The lower molecular weight and lower sulfur content, indicates that the sulfur linkages are being first attacked, resulting in ring opening which in turn weakens the asphaltene and brings about fragmentation. The remaining particle is more resistant to further reaction, requiring significantly higher temperature.

In a second series of operations, the crude tower bot toms was processed at 3,000 p.s.i.g., a liquid hourly space velocity of 1.0 (on fresh feed, 1.6 on the combined feed including recycle of a portion of the 444 C. drag stream from the total product) and a maximum catalyst temperature of 427 C. Analyses were performed on a portion of the drag stream, and the pertinent data are presented in the follow Table II, and identified as stream A. This drag stream was reprocessed at 3,000 p.s.i.g., a liquid hourly space velocity of 2.0 and a temperature gradient of 420 C. to 450 C. The 444 C.-plus portion of the product is identified in Table II as stream B. A portion of stream B was again processed at 3,000 p.s.i.g., a liquid hourly space velocity of 2.0 and a temperature gradient of 420 C. to 446 C. The 444 C.-plus portion of this product is identified in Table II as stream C.

TABLE II.-444 C., PLUS PRODUCT ANALYSES It should be noted that the residual asphaltenes in stream A, having a hydrogenzcarbon ratio of 0.99, are not structurally dissimilar from the native asphaltenes in the charge stock. This is not the case with respect to stream B where the residual asphaltenes have a hydrogenzcarbon ratio of 0.86. Furthermore, it should be noted that the asphaltenes in stream B appear to be resistant to further conversion as indicated by stream C analyses.

Streams A and C were then individually and separately subjected to further conversion in an autoclave operation at a charge to catalyst weight ratio of 10:1. The catalyst was 2.0% nickel and 16.0% molybdenum (by weight, and computed as the elemental metal) combined with a carrier material of 68.0% alumina, 22.0% boron phosphate and 10.0% silica. 200 grams of the stream and 20 grams of the catalyst Were pressured to atmospheres of hydrogen and heated to a temperature of 400 C., the final pressure being 200 atmospheres. After a period of four (4) hours, the autoclave was cooled, depressured,

TABLE III.--444 C., PLUS PRODUCT ANALYSES Stream Designation API at 60 F 25. l 19. 6 Heptane-insolubles, wt. percent. 0. 15 O. 89 Asphaltene conversion, wt. percent 87. 43. 1

The difference in the degrees to which the residual asphaltenes are prone to further conversion is evident from the data presented in Table III. In the case of the asphaltenes originally catalytically processed at below 430 C., there has been an ultimate conversion of 87.5% compared to only 43.1% where 'the asphaltenes had originally been subjected to temperatures above 430 C.

Utilization of the foregoing concepts is best illustrated with respect to a commercially scaled unit of about 5,000 bbl./day capacity. The charge stock is a reduced Kuwait crude having a gravity of 144 API at 60 F., and con taminated by 4.1% by weight of sulfur, 2,140 ppm. of nitrogen, about 55 p.p.m. of nickel and vanadium porphyrins and about 3.0% by weight of heptane-insolubles. The initial boiling point is 640 F., the 50% distillation point is 985 F. and 45% by volume boils above a temperature of 1,025 F. The object is to produce the maximum quantity of middle-distillates, principally a kerosene fraction boiling from 330 F. to 500 F. and a gas oil fraction boiling from 500 F. to 700 F.

The charge stock is initially admixed with 10,000 s.c.f./ bbl. of hydrogen (inclusive of 1,800 s.c.f./bbl. of makeup hydrogen consumed in the overall process) and a normally liquid recycle stream having an initial boiling point of 850 F. The recycle stream is in an amount such that the combined feed ratio of normally liquid hydrocarbons is 2.0, or about 5,000 bbL/day, the fresh feed liquid hourly space velocity being 0.5. The catalytic composite contains 1.8% by weight of nickel and 16.0% by weight of molybdenum combined with an alumina-silica-boron phosphate carrier material. The pressure on the reaction zone is about 2,750 p.s.i.g. and the catalyst bed inlet temperature is controlled such that the maximum catalyst temperature in the reaction zone does not exceed 400 C. (752 F.). The product effluent is fractionated at a cut-point of 454 C. (850 F.) to provide a bottoms fraction in an amount of about 7,000 bbl./ day. Of this heavier fraction, 5,000 bbL/day are recycled to combine with the fresh charge stock as hereinabove set forth.

The remaining portion of the heavier fraction, about 2,000 bbl./day, is passed into a second reaction zone with about 10,000 s.c.f./bbl. of hydrogen at a pressure of about 2,600 p.s.i.g. The catalyst is identical to that employed in the first reaction zone, and is in an amount such that the liquid hourly space velocity is about 0.75. The temperature at the catalyst bed inlet is controlled such that the maximum temperature is about 420 C. (788 F). The effluent from the second zone is flashed in a vacuum column functioning at a pressure of 55 mm. of Hg absolute, and at substantially the same temperature as the product eflluent. A residuum fraction is removed from the vacuum column in an amount of 420 bbl./day, or about 8.4 volume percent of the fresh charge stock.

The flash material is combined with the lighter fraction from the fractionator and about 8,000 s.c.f./bbl. of hydrogen. The mixture is heated to a temperature of about 440 C. (824 F.) and is introduced into a reaction zone at a pressure of about 2,500 p.s.i.g. The catalytic composite within this third reaction zone consists of 1.8% nickel and 16.0% molybdenum combined with a carrier material of 63.0% alumina and 37.0% silica, and the feed mixture contacts the catalyst at a liquid hourly space velocity of 1.0. The temperature differential across the catalyst bed is such that the product effluent is at a temperature of about 470 C. (878 F.).

The product effluent is passed into a high pressure separator functioning at a temperature of about F., and from which a hydrogen-rich gaseous phase is withdrawn by compressive means and recycled to the three reaction zones. The normally liquid stream from the high pressure separator is fractionated into the desired boiling ranges following the removal of dissolved hydrogen and light gaseous hydrocarbons. The following Table IV presents the overall yields including the residuum fraction withdrawn from the vacuum column. All the liquid streams are clean, being substantially free from both sulfur and nitrogen.

TABLE IV.OVERALL PRODUCT YIELDS Wt. Vol. BbL/day percent percent Component:

Hydrogen (2. 79) (1, 788) Arnlnonia 0.20 Hydrogen sulfide 3. 89 ethane 0.67 Ethane 0.83 Propane 1. 51 Butane 104 2. 32 Pentane 169 2. 19 Hexane 225 3. 23 Heptane, 330 867 13. 31 330 F.500 F. 1, 542 26.12 500 F.700 F. 2, 226 39. 07 Reslduurn 420 9. 45

1 Hydrogen consumed, wt. percent and s.c.f./bb1.

The foregoing specification, particularly the examples thereof, indicates the method by which my invention is effected, and clearly shows the benefits afforded the use thereof as an integral part of a process for converting asphaltene-containing hydrocarbon mixtures.

I claim as my invention:

1. A process for the conversion of an asphaltene-containing hydrocarbon charge stock which comprises the steps of:

(a) admixing said charge stock with hydrogen and catalytically reacting the mixture in a first reaction zone at a maximum catalyst temperature below about 430 C., separating the resulting first reaction zone eflluent to provide a first fraction having an end boiling point of about 440 C. to about 460 C., and a concentrated unconverted asphaltene fraction;

(b) reacting at least a portion of said fraction containing unconverted asphaltenes in a second catalytic reaction zone at a maximum catalyst temperature below about 430 C., recycling the remaining portion of said unconverted asphaltene-containing fraction to combine with said charge stock;

(c) separating the resulting second zone effluent to provide a second fraction and to concentrate unconverted asphaltenes in a residuum fraction;

((1) combining said second fraction with at least a portion of said first fraction; and,

(e) catalytically reacting the resulting mixture in a third reaction zone at a catalyst temperature of about 43 0 C. to about 510 C.

2. The process of claim 1 further characterized in that said maximum catalyst temperature within the first and second reaction zone is within the range of 320 C. to 430 C.

3. The process of claim 1 further characterized in that said first reaction zone effluent is separated to provide a first fraction containing hydrocarbons boiling below 1 l 1 2 about 454 C., and a second fraction boiling above 454 FOREIGN PATENTS C. and containing the unconverted asphaltenes. 502,798 3/1939 Great Britain References Cited DELBERT E. GANTZ, Primary Examiner UNITED STATES PATENTS 5 T. H. YOUNG, Assistant Examiner 2,526,966 10/1950 Oberfell et a1. 20857 3,055,822 9/1962 Honerkamp et a1. 20861 US. Cl. X.R.

3,409,538 11/1968 Gleim et a1. 20859 20857, 59 

